In many catalytic reactions, it is important that reactants be well mixed with catalyst to afford sufficient opportunity for the reactant to contact the catalyst. Fluidized reactors have been designed to ensure adequate mixing of catalyst and reactants. Fluidized reactors are designed to ensure that the reactants are in contact with the catalyst for sufficient time to allow for the reaction to proceed. However, in many catalyzed reactions the reactants should not remain in contact with catalyst too long or overconversion can occur which can generate undesirable byproducts and degrade product quality. This is especially true when the reaction involves hydrocarbons of which overreaction can cause excess generation of coke, inhibiting catalyst activity and selectivity.
Space velocity typically referred to as weight hourly space velocity (WHSV) is crucial to ensuring that reactants and catalyst are in contact for the optimal duration. Space velocity is a reaction condition that is important when rapid reaction times are involved such as monomolecular catalytic cracking reactions or rapid catalytic conversion reactions. The catalyst and reactants need to make contact, but excessive contact time will cause additional undesirable reaction to occur. Space velocity is calculated by Formula 1:
                    WHSV        =                              M            f                                m            c                                              (        1        )            where WHSV is the weight hourly space velocity, Mf is the mass flow rate of feed to the reactor and mc is the mass of catalyst in the reactor. Space time is the inverse of space velocity. The mass of catalyst, mc, can be determined by Formula 2:
                              m          c                =                              Δ            ⁢                                                  ⁢                          P              ·              V                                h                                    (        2        )            where ΔP is the pressure drop over the height, h, of the reactor and V is the volume of the reactor. The ratio of pressure drop and height is the catalyst density in the reactor:
                              ρ          c                =                              Δ            ⁢                                                  ⁢            P                    h                                    (        3        )            
Hence, combining Formulas 2 and 3:mc=ρc·V  (4)Accordingly, both density, ρc, and space velocity, WHSV, are functions of pressure drop, ΔP. From Formula 4, the relationship of space velocity and volume is shown in Formula 5:
                    WHSV        =                              M            f                                              ρ              c                        ⁢            V                                              (        5        )            
Catalyst flux is determined by Formula 6:
                              ω          c                =                              M            c                    A                                    (        6        )            where ωc is catalyst flux, Mc is the mass flow rate of catalyst and A is the cross sectional area of the reactor. Additionally the product of height, h, and cross-sectional area, A, when constant, of the reactor is the reactor volume, V:V=h·A  (7)
The mass flow rate of feed, Mf, to the reactor is calculated by Formula 8:Mf=vf·ρf·A  (8)where vf is the superficial gas velocity of the feed, ρf is the density of the feed and A is the cross sectional area of the reactor at which the velocity is measured. Hence, substituting Formulas 2, 7 and 8 into Formula 1 for constant cross-sectional reactor area yields formula 9:
                    WHSV        =                                            v              f                        ⁢                          ρ              f                                            Δ            ⁢                                                  ⁢            P                                              (        9        )            
Residence time is a reaction condition that is important when the reaction is not as rapid. The catalyst and reactants need to soak together to ensure catalyst and reactants are in contact and for a sufficient period of time to allow the reaction to occur. Residence time, Tr, is calculated by Formula 10:
                              T          r                =                  V                      Q            f                                              (        10        )            where Qf is the actual volumetric flow rate of feed at reactor process conditions of temperature and pressure. The volumetric flow rate, Qf, to the reactor is calculated by Formula 11:Qf=vf·A  (11)Substituting Formulas 7 and 11 into Formula 10 for constant cross-sectional reactor area yields:
                              T          r                =                  h                      v            f                                              (        12        )            In a fluidized catalytic reactor, the flow characteristics may be considered to assure space velocity or residence time is optimal.
Two types of fluidization regimes typically used in fluidized catalytic reactors are a transport flow regime and a bubbling bed. Transport flow regimes are typically used in FCC riser reactors. In transport flow, the difference in the velocity of the gas and the catalyst, called the slip velocity, is relatively low, typically less than 0.3 m/s (1.0 ft/s) with little catalyst back mixing or hold up. Slip velocity is calculated by formula 9:
                              v          s                =                                            v              f                        ɛ                    -                      v            c                                              (        13        )            where vs is the slip velocity, vf is the superficial gas velocity of the feed, vc is the catalyst velocity and ε is the void fraction of the catalyst. Another way to characterize flow regimes is by slip ratio which is the ratio of actual density in the flow zone to the non-slip catalyst density in the flow zone. The non-slip catalyst density is calculated by the ratio of catalyst flux to the superficial gas velocity as in formula 10:
                              ρ          cns                =                              ω            c                                v            f                                              (        14        )            where ρcns is the non-slip catalyst density in the flow zone, ωcc flux of the catalyst and vf is the superficial gas velocity of the feed. The slip ratio is proportional to the hold up of catalyst in the flow zone. Typically, a slip ratio for a transport flow regime does not reach 2.5. Consequently, the catalyst in the reaction zone maintains flow at a low density and very dilute phase conditions. The superficial gas velocity in transport flow is typically greater than 3.7 m/s (12.0 ft/s), and the density of the catalyst is typically no more than 48 kg/m3 (3 lb/ft3) depending on the characteristics and flow rate of the catalyst and vapor. In transport mode, the catalyst-vapor mixture is homogeneous without vapor voids or bubbles forming in the catalyst phase.
Fluidized bubbling bed catalytic reactors are also known. In a bubbling bed, fluidizing vapor forms bubbles that ascend through a discernible top surface of a dense catalyst bed. Only catalyst entrained in the vapor exits the reactor with the vapor. The superficial velocity of the vapor is typically less than 0.5 m/s (1.5 ft/s) and the density of the dense bed is typically greater than 480 kg/m3 (30 lb/ft3) depending on the characteristics of the catalyst. The mixture of catalyst and vapor is heterogeneous with pervasive vapor bypassing of catalyst.
Intermediate of dense, bubbling beds and dilute, transport flow regimes are turbulent beds and fast fluidized regimes. U.S. Pat. No. 4,547,616 discloses a turbulent bed used in a reactor for converting oxygenates to olefins. In a turbulent bed, the mixture of catalyst and vapor is not homogeneous. The turbulent bed is a dense catalyst bed with elongated voids of vapor forming within the catalyst phase and a less discernible surface. Entrained catalyst leaves the bed with the vapor, and the catalyst density is not quite proportional to its elevation within the reactor. The superficial velocity is between about 0.5 and about 1.3 m/s (1.5 and 4.0 ft/s), and the density is typically between about 320 and about 480 kg/m3 (20 and 30 lb/ft3) in a turbulent bed.
U.S. Pat. No. 6,166,282 discloses a fast fluidized flow regime for oxygenate conversion. Fast fluidization defines a condition of fluidized solid particles lying between the turbulent bed of particles and complete particle transport mode. A fast fluidized condition is characterized by a fluidizing gas velocity higher than that of a dense phase turbulent bed, resulting in a lower catalyst density and vigorous solid/gas contacting. In a fast fluidized zone, there is a net transport of catalyst caused by the upward flow of fluidizing gas. The superficial combustion gas velocity for a fast fluidized flow regime is conventionally believed to be between about 1.1 and about 2.1 m/s (3.5 and 7 ft/s) and the density is typically between about 48 and about 320 kg/m3 (3 and 20 lb/ft3). Catalyst exits the reaction zone a small amount slower than the vapor exiting the reaction zone. Hence, for a fast fluidized flow regime the slip velocity is typically greater than or equal to 0.3 m/s (1.0 ft/s) and the slip ratio is greater than or equal to 2.5 for most FCC catalysts. Fast fluidized regimes have been used in FCC combustors for regenerating catalyst and in coal gasification.
The conversion of hydrocarbon oxygenates to olefinic hydrocarbon mixtures is accomplished in a fluidized catalytic reactor. Such oxygenates to olefins reactions are rapidly catalyzed by molecular sieves such as a microporous crystalline zeolite and non-zeolitic catalysts, particularly silicoaluminophosphates (SAPO). Numerous patents describe this process for various types of these catalysts: U.S. Pat. Nos. 3,928,483; 4,025,575; 4,252,479; 4,496,786; 4,547,616; 4,677,243; 4,843,183; 4,499,314; 4,447,669; 5,095,163; 5,191,141; 5,126,308; 4,973,792 and 4,861,938.
The oxygenates to olefins catalytic process may be generally conducted in the presence of one or more diluents which may be present in the hydrocarbon oxygenate feed in an amount between about 1 and about 99 mol-%, based on the total number of moles of all feed and diluent components fed to the reaction zone (or catalyst). Diluents include, but are not limited to, helium, argon, nitrogen, carbon monoxide, carbon dioxide, hydrogen, water, and hydrocarbons such as methane, paraffins, aromatic compounds, or mixtures thereof. U.S. Pat. Nos. 4,861,938 and 4,677,242 particularly emphasize the use of a diluent combined with the feed to the reaction zone to maintain sufficient catalyst selectivity toward the production of light olefin products, particularly ethylene.
U.S. Pat. No. 6,023,005 discloses a method for selectively converting hydrocarbon oxygenates to light olefins in which desirable carbonaceous deposits are maintained on the total reaction volume of catalyst by regenerating only a portion of the total reaction volume of catalyst and mixing the regenerated portion with the unregenerated total reaction volume of catalyst. The method incorporates a fluidized catalytic bed reactor with continuous regeneration. In a preferred arrangement, the oxygenate feed is mixed with regenerated catalyst and coked catalyst at the bottom of a riser and the mixture is lifted to a disengaging zone. In the disengaging zone, coked catalyst is separated from the gaseous materials by means of gravity or cyclone separators. A portion of the coked catalyst to be regenerated is sent to a stripping zone to recover adsorbed hydrocarbons. Stripped spent catalyst is passed to a regenerator.
U.S. Pat. No. 4,547,616 discloses an improvement in a process for the conversion of hydrocarbon oxygenates to olefins by the operation of a turbulent bed at elevated temperatures and controlled catalyst activity. The turbulent bed is maintained in a vertical reactor column to achieve good mixing at a velocity greater than the dense bed transition velocity to a turbulent regime and less than transport velocity for the average catalyst particle. The superficial fluid velocity is disclosed in a range between about 0.3 to 2 meters per second. Provision is made for passing partially regenerated catalyst to the reactor fluidized bed of catalyst beneath the upper interface and sufficiently below to achieve good mixing in the fluid bed.
Another typical fluidized catalytic reaction is a fluidized catalytic cracking (FCC) process. An FCC process is carried out by contacting the starting material whether it be vacuum gas oil, reduced crude, or another source of relatively high boiling hydrocarbons with a catalyst made up of finely divided or particulate solid material. The catalyst is fluidly transported by passing gas through it at sufficient velocity to produce a transport flow regime. Contact of the oil with the fluidized catalytic material catalyzes the cracking reaction. The cracking reaction deposits coke on the catalyst. Catalyst exiting the reaction zone is spoken of as being “spent”, i.e., partially deactivated by the deposition of coke upon the catalyst. Coke is comprised of hydrogen and carbon and can include other materials in trace quantities such as sulfur and metals that enter the process with the starting material. Coke interferes with the catalytic activity of the spent catalyst by blocking acid sites on the catalyst surface where the cracking reactions take place. Spent catalyst is traditionally transferred to a stripper that removes adsorbed hydrocarbons and gases from catalyst and then to a regenerator for purposes of removing the coke by oxidation with an oxygen-containing gas. The regenerator may operate with a bubbling bed, turbulent bed or fast fluidized flow regime. Such regenerators using a fast flow regime are called combustors. However, in a regenerator or combustor, coke is burned from the catalyst. The catalyst does not provide a catalytic function other than with regard to oxidation. An inventory of catalyst having a reduced coke content, relative to the spent catalyst in the stripper, hereinafter referred to as regenerated catalyst, is collected for return to the reaction zone. Oxidizing the coke from the catalyst surface releases a large amount of heat, a portion of which escapes the regenerator with gaseous products of coke oxidation generally referred to as flue gas. The balance of the heat leaves the regenerator with the regenerated catalyst. The fluidized catalyst is continuously circulated between the reaction zone and the regeneration zone. The fluidized catalyst, as well as providing a catalytic function in the reaction zone, acts as a vehicle for the transfer of heat from zone to zone. The FCC processes, as well as separation devices used therein are fully described in U.S. Pat. Nos. 5,584,985 and 4,792,437. Specific details of the various reaction zones, regeneration zones, and stripping zones along with arrangements for conveying the catalyst between the various zones are well known to those skilled in the art.
The FCC reactor catalytically and thermally cracks gas oil or heavier feeds into a broad range of products. Cracked vapors from the FCC unit enter a separation zone, typically in the form of a main column, that provides a gas stream, a gasoline cut, light cycle oil (LCO) and clarified oil (CO) which includes heavy cycle oil (HCO) components. The gas stream may include dry gas, i.e., hydrogen and C1 and C2 hydrocarbons, and liquefied petroleum gas (“LPG”), i.e., C3 and C4 hydrocarbons. The gasoline cut may include light, medium and heavy gasoline components. A major component of the heavy gasoline fraction comprises condensed single ring aromatics. A major component of LCO is condensed bicyclic ring aromatics.
Subjecting product fractions to additional reactions is useful for upgrading FCC product quality. The recracking of heavy product fractions from an initially cracked FCC product is one example. Typically, in recracking, cracked effluent from a riser of an FCC reactor is recontacted with catalyst at a second location to cleave larger molecules down into smaller molecules. For example, U.S. Pat. No. 4,051,013 discloses cracking both gasoline-range feed and gas oil feed in the same riser at different elevations. Such reactions are relatively rapid. WO 01/00750 A1 discloses introducing gasoline feed and FCC feed at different elevations in a riser reactor, separating the cracked product and recycling portions thereof back to the same riser reactor. Other types of reactions to upgrade FCC product fractions are less rapid.
The above-described hydrocarbon catalytic conversion processes are sensitive to underreaction and overreaction which both degrade product quality. Use of a fast fluidized flow regime assures thorough mixing of catalyst and feed to catalyze the reaction. Hence, we sought to provide an improved fluidized non-oxidative catalytic hydrocarbon conversion process and apparatus that can provide a fast fluidized flow regime at adjustable flow conditions that will enhance the conversion to the desired products. Additionally, we sought to provide a reactor that can accommodate the varying demands on space velocity and residence time based on different feed composition and desired products.